Method to extend the utilization of a catalyst in a multistage reactor system

ABSTRACT

The invention provides a method to extend the utilization of a catalyst in a multistage reaction system, provided that a primary reaction and a secondary (or more) catalyst-poisoning reaction occur on the same catalyst, and the rate of the secondary (or more), catalyst-poisoning reaction is faster than the rate of the primary reaction.

CROSS REFERENCE TO RELATED APPLICATION

This application claims the benefit of U.S. Provisional Application No. 60/615,083, filed Oct. 1, 2004, which is hereby incorporated by reference.

FIELD OF THE INVENTION

This invention relates to a method to extend utilization of a catalyst in a multistage reactor system. More specifically, the invention relates to a method to extend catalyst utilization in sets of reactions catalyzed on the same catalyst (i.e., hydrogenation and desulfurization), where one reaction causes catalyst poisoning.

BACKGROUND

Hydrogenation is a commonly practiced process in petroleum, chemical, and food industries. Depending on the feedstocks and the severity of the operating conditions (e.g., temperature, pressure, and contact time), the hydrogenation process may saturate unsaturated bonds, reduce aldehydes and ketones, reduce carboxylic acids and their esters, reduce nitrogen-containing compounds, reduce sulfur-containing compounds, and cause numerous other reduction reactions.

In catalytic hydrogenation processes, the active catalyst is usually a zero-valent metallic catalyst of one or more of the Group VIII elements of the periodic table. Catalyst deactivation is a common phenomenon in catalytic processes (e.g., hydrocracking, hydrotreating, as well as hydrogenation). For example, sulfur in a feedstock is a poison for most hydrogenation catalysts. In such cases, hydrodesulfurization proceeds along with hydrogenation. Reduction of the organic sulfur compounds in the feedstock converts the sulfur to its divalent form. The divalent sulfur reacts with the active zero-valent metal of the catalyst to form a metal sulfide. At low to moderate temperatures (lower than about 200° C.) of hydrogenation, metal sulfides do not have practical hydrogenation activity and thus are considered the poisoned form of the metallic catalysts.

To preclude or reduce catalyst poisoning, desulfurization of the feedstock prior to the desired reaction can be achieved by physical adsorption of the organoo-sulfur compounds on certain reactive adsorbents. “Reactive adsorption” refers to cases where the sulfur compounds react with the adsorbent. Various oxides (including ZnO, CuO, and MnO) can react with the sulfur compounds (specifically hydrogen sulfide) and remove them from gaseous streams.

At temperatures above 200° C., metal sulfides show hydrodesulfurization and some hydrogenation activity. Hydrodesulfurization is a widely practiced process in the petroleum industry to selectively desulfurize feedstocks. In hydrodesulfurization, the organo-sulfur compounds are converted to hydrogen sulfide, which is usually removed from the reactor as a gas mixed with the excess hydrogen. A thorough review of desulfurization processes has recently been published by I. V. Babich and J. A. Moulijn (Fuel, 82, pp 607-631, (2003)).

In a single reactor, the poisoned zone will progress downstream, gradually diminishing the activity of the upstream portion of the catalyst bed to the point where the reactor cannot produce a product with desired specifications. Although the entire reactor bed cannot meet the required specification, the catalyst downstream from the poisoned zone still is relatively active and may only be slightly deactivated.

In traditional practice, once the quality of the reactor product falls below the accepted specification, it becomes necessary to replace the catalyst of the entire reactor bed. This requires stopping the operation to remove and dispose of the poisoned catalyst once the quality of the reactor product falls below the accepted specification, and to load the reactor with fresh catalyst before resuming operation. Because the rate of the secondary reaction (for example, hydrodesulfurization) is faster than the rate of the primary reaction (for example hydrogenation), the full extent of the useful life of the catalyst is not realized and the economics of the process are negatively affected.

Attempts have been made to extend catalyst utilization. These methods include external or in-situ regeneration, and various process modifications, including altering operating conditions (such as increasing temperatures) to compensate for the lost primary reaction activity, for example, a process having a multi-reactor or a single reactor with multi-zone concept for selective hydrogenation of diolefins in a liquid hydrocarbon. Once the zone or reactor furthest from the inflow of feedstock loses its activity, the next closest zone or reactor is brought on stream and placed upstream of the partially deactivated zone.

The problem that remains to be solved is how to extend catalyst utilization when a primary reaction (e.g., hydrogenation) and a secondary (or more) catalyst-poisoning reaction (e.g., hydrodesulfurization) occur on the same catalyst, and when the rate of the poisoning reaction is faster than the rate of the primary reaction.

SUMMARY OF THE INVENTION

A method is provided for extending catalyst utilization in a multistage reaction system comprising: a) passing a feedstock and hydrogen through at least two serially-connected reactors in positions R1, R2, . . . Rn, wherein n is the number of reactors, each reactor containing a catalyst, for a period until catalyst in at least one most upstream reactor is deactivated for a secondary catalytic reaction or until the product from the most downstream reactor fails to meet a desired specification; b) bypassing the at least one most upstream reactor of step (a) to pass the feedstock and hydrogen into at least one downstream reactor; c) reloading the at least one bypassed reactor of step (b) with fresh catalyst; d) placing the at least one reloaded reactor of step (c) downstream of at least one of the serially-connected reactors that were not reloaded with fresh catalyst in step (c); and e) repeating steps (a) through (d) to meet the product specification.

In one embodiment of the invention, the temperature may be optionally increased after step (a) to continuously meet the specification for the product as it leaves the most downstream reactor. Subsequently, after step (d) the temperature is decreased to its original level. Adjusting the temperature allows the product specification to be met continuously even during catalyst change out without interrupting the process.

The method of the invention uses a catalyst selected from the group consisting of a zero-valent element of one or more of the Group VIII elements of the Periodic Table. One embodiment of the method for extending catalyst utilization provides: a) passing an organic feedstock and hydrogen through at least two serially-connected reactors in positions, R1, R2, . . . Rn, wherein n is the number of reactors, each reactor containing a porous metal or supported catalyst, for a period until the catalyst in at least one most upstream reactor is deactivated for desulfurization or until the product from the most downstream reactor fails to meet a desired specification; b) bypassing the at least one most upstream deactivated reactor of step a) to pass the feedstock and hydrogen into downstream reactors; c) reloading the at least one bypassed reactor of step b) with fresh catalyst; d) placing the at least one reloaded reactor of step c) downstream; and e) repeating steps (a) through (d) to meet the product specification.

Another embodiment of the invention extends catalyst utilization in the manufacture of biologically derived 1,3-propanediol comprising the steps: a) passing a biologically derived organic feedstock comprising 1,3-propanediol and hydrogen through at least two serially-connected reactors in positions R1, R2, . . . . Rn, wherein n is the number of reactors, each reactor containing a catalyst comprising nickel with heavy nickel loading supported on extrudates of silica/alumina, for a period until the catalyst in the most upstream reactor is deactivated for desulfurization or until the product from the most downstream reactor fails to meet a desired specification; b) bypassing the reactor in position R1 to pass the feedstock and hydrogen into the reactor in position R2; c) reloading the bypassed reactor of step b) with fresh catalyst; d) placing the reloaded reactor of step c) in the reactor series downstream at position Rn; and e) repeating steps (a) through (d) as necessary to meet the product specification.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 a shows the initial arrangement of reactors in series where the feedstock enters reactor A (in position R1), and then downstream reactors B and C (in positions R2, . . . Rn), with optional intermittent hydrogen input.

FIG. 1 b shows a flow diagram for the subsequent arrangement of the reactors after the deactivated catalyst in reactor A is replaced with fresh catalyst and is placed back on stream in position Rn, where the outlet of reactor C supplies the feed for reactor A.

FIG. 2 shows the distribution of sulfur on two spent catalyst beds.

FIG. 3 shows the activity of fresh and spent catalysts of varying sulfur content. Catalyst activity (expressed as a percent reduction in absorbance at UV-270 nm) clearly decreases with increased sulfur deposited on the catalyst.

FIG. 4 shows the effect of temperature on the profile of sulfur deposited on the spent catalyst. At 120° C., most of the sulfur accumulates near the entrance of the reactor, whereas at 80° C., sulfur deposition is spread more uniformly throughout the catalyst bed.

FIG. 5 shows that a reactor filled with partially-poisoned catalyst can remove most of the sulfur and serve as a guard bed to protect downstream reactor beds from heavy sulfur deposition load.

DETAILED DESCRIPTION OF THE INVENTION

The inventors have solved the stated problem with a method for extending catalyst utilization in a multistage reactor system for feedstocks containing organics. When a primary reaction (e.g., hydrogenation) and a secondary (or more) catalyst-poisoning reaction (e.g., hydrodesulfurization) occur on the same catalyst, and when the rate of the catalyst-poisoning reaction is faster than the rate of the primary reaction, catalyst poisoning is mainly restricted to the upstream reactor while the downstream reactor completes the bulk of hydrogenation needed to meet the product specification. By “reactor” is meant an individual reactor within a series of multiple reactors.

The invention directs the catalyst-poisoning reaction to the upstream reactor, preventing the downstream reactors from uniform and/or rapid poisoning. Because the upstream reactor is used as a guard bed to protect the catalyst in the downstream reactor or reactors against poisoning, the overall useful life of the catalyst is extended.

This method provides two significant advantages for multistage reaction systems: 1) catalyst utilization can be significantly extended, and 2) process and equipment downtime needed to change the catalyst can be significantly reduced or even eliminated.

The method uses a multistage reactor system comprised of a minimum of two reactors in series (in positions R1, . . . Rn, where n is the number of reactors). (See FIG. 1 a.) The reactors are operated under such conditions as to concentrate or direct the secondary (or more) catalyst-poisoning reaction to the most upstream reactor(s). Secondary (or more) reactions may be, for example, desulfurization or demetallation. When the most upstream reactor in position R1 is poisoned, it is taken out of service (bypassed) and its catalyst is changed to fresh catalyst. During this period, the process continues in the second or downstream reactors (in positions R2, . . . Rn). To compensate for the removed or bypassed reactor, the temperature in the reactor now in the most upstream position R1 can be temporarily increased to meet the product specification requirements. After reloading the bypassed reactor with fresh catalyst, it is placed downstream in the train of reactors (in one embodiment at the most downstream position, Rn) and the operating temperatures can be adjusted to reflect the number of reactors in the train. The reloaded reactor in position Rn now primarily serves as the reactor for the primary reaction (e.g., hydrogenation). As shown in FIG. 1 b, the reactor now most upstream in position R1 (containing partially-poisoned catalyst) serves as the primary site for the secondary reaction, thus protecting downstream reactors in positions R2, . . . Rn. The cycling of the reactors continues as the next reactor most upstream in the process train deactivates to a designated level, is removed or bypassed, refreshed, and replaced downstream in turn.

Catalyst

Any hydrogenation catalyst known in the art is suitable for use in this invention, provided that the primary hydrogenation reaction and a secondary (or more) catalyst-poisoning reaction occur on the same catalyst, and the rate of the catalyst-poisoning reaction is faster than the rate of the primary hydrogenation reaction. By varying the operating conditions (e.g. temperature) the rate of hydrogenation relative to the rate of one or more poisoning reactions can be adjusted to remove the bulk of the poisons in the upstream reactor. In one embodiment of the invention, the hydrogenation catalyst can remove sulfur in a narrow band. In this case, the catalyst's ability to remove sulfur in a narrow band is used to vary the relative rate of hydrogenation relative to the rate of desulfurization and the affinity of the catalyst to react with the sulfur to remove it. The relative rates of the reactions in each set can be determined by one of ordinary skill in the art by varying the reaction temperature, pressure, and the feedstock contact time with the catalyst without generating undesirable side reactions.

The catalyst used in the invention is comprised of at least one zero-valent element of the Group VIII elements of the Periodic Table. In embodiments of the invention the catalyst is at least one of Fe, Co, Ni, Ru, Rh, Pd, Os, Ir, and Pt, with or without various promoters. The catalyst need not be present in its elemental form. The promoter may be any element of the periodic table or a compound thereof that could be added to the catalyst to enhance its activity or selectivity. “The Handbook of Heterogeneous Catalytic Hydrogenation for Organic Synthesis,” by Shigeo Nishimuru, John Wiley (2001), ISBN: 0471-39698-2, extensively discusses such catalysts and conditions.

The catalyst may be a porous metal structure, a Raney® catalyst, or supported on a substrate. The catalyst support could be from any support material known in the art, such as at least one of carbon, alumina, silica, titania, silica-alumina, silica-titania, titania-alumina, clays, aluminosilicates, zeolites, water insoluble salts of calcium (such as calcium carbonate), barium (such as barium sulfate), strontium (such as strontium carbonate), and compounds and combinations thereof. The catalyst may have various shapes or sizes, such as fine powder, granules, tablets, pellets, extrudates, or other structured supports.

Feedstocks

Suitable feedstocks for the process of this invention are those comprising compounds that can be hydrogenated, as well as other materials or compounds that poison the catalyst in a secondary reaction. The invention is useful for hydrogenation/desulfurization systems in which the feedstocks contain sulfur compounds. The sulfur-containing feedstock is not limited to petroleum-based hydrocarbons and may be any organic fluid/s derived from fossil and/or biological sources. The inventive technique is more advantageously applied to feedstocks characterized by low but still undesirable levels of sulfur. The lower the undesirable sulfur level in the feedstock, the longer the life of the catalyst can be extended by use of the invention.

Operating Conditions

The operating conditions for the invention are first selected to meet the specification for the product produced by the particular catalytic reactions. Those skilled in the art will be well aware of the methods to adapt the invention to yield a particular product, for instance, conducting a series of experiments to determine the best temperature, pressure, and the feedstock contact time to concentrate the catalyst-poisoning reaction in the upstream reactor(s). Additionally, it will be further beneficial to test the activity of the partially poisoned catalyst from the downstream reactor to demonstrate and measure its remaining activity available for the primary hydrogenation reaction. Once these two conditions are established and verified, a train of multiple reactors can be designed to implement this invention for a particular product.

The claimed invention describes a multistage reactor system to manage catalyst poisoning from the secondary reaction (i.e., hydrodesulfurization) in a more economical way than previously known. Instead of one large reactor typically used in conventional systems, at least two smaller reactors having a combined catalyst volume equivalent to that of the one large reactor are placed in series. The reactors are used as disclosed to hydrogenate the feedstock until catalyst is poisoned for the secondary reaction or the product fails to meet the product specification as it exits the most downstream reactor. The first reactor(s) is then bypassed and the feedstock is directed instead into the next reactor(s) downstream in the series.

Increasing the processing temperature to compensate for the removal or bypassing of the upstream reactor enables the remaining reactor(s) to continuously meet the product specification during catalyst change out without any interruption in the process. During this operating period, the first reactor(s) is reloaded with fresh catalyst. The renewed reactor(s) is/are then returned to service, but placed downstream of the partially deactivated reactor(s), preferably in the most downstream position in the train and the temperature is adjusted to the desired original level such that the product specifications are met. The reactor(s) containing partially poisoned catalyst, now at the most upstream position in the series, act(s) as the primary desulfurization reactor, while the renewed reactor(s) with the fresh catalyst provides primarily hydrogenation activity. A schematic of a multi-feed embodiment of the process is shown in FIGS. 1 a and 1 b. The hydrogen may be fed into individual reactors or may be fed to the most upstream reactor.

When the primary reaction is hydrogenolysis (such as hydrodesulfurization, hydrodenitrogenation or hydrodeoxygenation), the operating temperature is usually at or below 400° C. When the primary reaction is hydrogenation, the operating temperature of this invention is usually at or below 200° C. (preferably in the range of about minus 50° C. to about 140° C.). During hydrogenation, the removed sulfur may deposit as adsorbed elemental sulfur or as a reacted compound (usually as a metal sulfide). Hydrogen flow rate and hydrogen pressure must be maintained to deliver adequate hydrogen to the catalyst surface to accomplish the desired hydrogenation. In general, the hydrogen feed rate depends on the hydrogen demand of the process. The operating hydrogen pressure for the process of this invention is above 100 kPa with a preferred range of 800-4240 kPa. In the hydrogenation of 1,3-propanediol, the hydrogen to crude PDO feed ratio is above 0.5 scc H2/g PDO with a preferred range of 1-20 scc H2/g PDO.

The process of this invention can be applied to any multi-reaction system where a primary reaction and a secondary (or more) catalyst-poisoning reaction occur on the same catalyst, and when the rate of the catalyst-poisoning reaction is faster than the rate of the primary reaction. These reactions can be of any chemistry where catalyst deactivation can be narrowed to a band (or portion) of the bed and only the contents of that deactivated band can be changed while the remainder of the reaction bed continues functioning. Examples of such multi-reaction systems include hydrocracking, hydrotreatment, hydrodeoxygenation, hydrodenitrogenation, and hydrodesulfurization reactions, where the catalyst may be poisoned by secondary hydrodemetallation reactions.

Reactor System

The invention can be applied to any type of reactor in multistage configuration, provided that the reactor configuration allows for the determination of the operating conditions that will concentrate the catalyst-poisoning phenomenon into the first reactor. Such reactors include fixed-bed catalytic reactors with upflow or down-flow arrangement, where the hydrogen can be fed individually into each reactor or fed just into one reactor. The hydrogen may flow co-currently or counter-currently with the liquid feedstock. The reactors may alternatively be of slurry-type or fluidized-bed or any other reactor type known in the literature (see, for example, Perry's Chemical Engineer's Handbook, Sixth Edition, R. H. Perry and D. Green, Ed.). An industrially advantageous reactor uses a packed-bed of catalyst wherein the liquid and gas flow co-currently or counter-currently, in an up-flow or down-flow (trickle-bed) mode of operation.

1,3-Propanediol Processed in a Hydrogenation/Desulfurization System

A suitable feedstock processed in a hydrogenation/desulfurization system comprises 1,3-propanediol (also hereinafter termed “PDO”), a monomer useful in the production of a variety of polymers including polyesters, polyurethanes, polyethers, and cyclic compounds. Homo- and co-polyethers of polytrimethylene ether glycol are examples of such polymers. The polymers are ultimately used in various applications including fibers, films, etc.

PDO may be obtained from non-renewable resources, typically petrochemical products. Chemical routes to generate PDO include hydroformylation of ethylene oxide over a catalyst or hydration of acrolein. Both of these synthetic routes to PDO involve the intermediate synthesis of 3-hydroxypropionaldehyde. The 3-hydroxypropionaldehyde is reduced to PDO in a final catalytic hydrogenation step. Subsequent purification involves several processes, including vacuum distillation. Hereinafter, the PDO from chemical processes is termed “chemical PDO” or “chemically derived PDO”.

PDO is also derived from renewable resources, including glucose or glycerol from such sources as corn or other biomass. Such PDO is hereinafter referred to as “biochemical PDO”, “bio-PDO”, or “biochemically-derived PDO”. The technique is disclosed in several patents, including U.S. Pat. Nos. 5,633,362; 5,686,276; and 5,821,092, all of which are incorporated in their entirety by reference herein. The PDO formed via biochemical routes contains numerous organic compounds and several organic sulfur compounds in the parts-per-million (ppm) range.

Applicants' invention is usefully practiced where hydrogenation is used as a polishing step in the production of PDO, for instance to obtain PDO of suitable quality for polymer production. (See WO2004/101479, WO2004/101482, and WO2004/101468, all of which are incorporated in their entirety by reference herein.) In the bio-PDO process, hydrogenation comprises contacting biochemically-derived PDO with hydrogen in the presence of a hydrogenation catalyst. The catalyst in the polishing process serves two purposes: 1) to hydrogenate the color and color precursor compounds, and 2) to remove the sulfur from the feedstock. The extent of hydrogenation can be determined as a function of color, residual carbonyls, iodine number, and similar indicators known to those of skill in the art.

The catalyst used in the invention is comprised of at least one zero-valent element of Group VIII of the periodic table. In embodiments of the invention the catalyst is at least one of Fe, Co, Ni, Ru, Rh, Pd, Os, Ir, and Pt. Various mixed metal oxides such as mixed copper, chromium, and zinc oxides are also effective catalysts for color removal. An embodiment of the invention utilizes a nickel catalyst with heavy nickel loading. Another embodiment of the invention utilizes a catalyst comprising nickel supported on extrudates of silica/alumina.

In another embodiment of the invention, the catalyst may be present with at least one promoter. The promoter may be any element of the periodic table or compound thereof that could be added to the catalyst to enhance its activity or selectivity. In other embodiments of the invention promoters are iron, chromium, and molybdenum.

The sulfur compounds contained in the crude PDO are reduced in the hydrogenation process. The reduced sulfur may then react with the hydrogenation catalyst, poisoning its active sites. This poisoning of the catalyst for hydrogenation by the desulfurization of the feedstock represents a significant cost in the manufacture of PDO. Previously when the catalyst lost its color-removing capacity, it was replaced with fresh catalyst regardless of any remaining ability to hydrogenate or remove sulfur from the feedstock. The replacement protocol required equipment downtime. In terms of materials and time, replacement of the underutilized catalyst was relatively expensive in light of the overall process. Cycling the reactors in the manner of the invention extends the overall utilization of the catalyst and can reduce or eliminate equipment downtime.

The temperature for the process ranges from about minus 50° C. to about 200° C. In another embodiment, the temperature for the process ranges from about 80° C. to about 140° C.

Other Sets of Multiple Reactions on the Same Catalyst

In petroleum feedstocks containing vanadium and/or nickel porphyrins, hydrotreatment and hydrodemetallation reactions are catalyzed on the same catalyst. The secondary reaction removes these metals and deposits them on the catalyst, poisoning the catalyst for the primary reaction. Similarly, in the upgrading of coal-derived liquids, the organometallic compounds of titanium and other elements present in the coal liquids are converted to metals by the secondary reaction, poisoning the catalyst for use in the primary reaction. Other sets of reactions that are catalyzed on the same catalyst are known to those of skill in the art of industrial catalysis.

Usually hydrodemetallation reactions are relatively faster than the hydrotreatment reactions on the same catalyst. By determining and selecting a set of suitable operating conditions, the deposition of metals can be concentrated in the most upstream reactor(s) of a multiple stage reactor system, allowing the downstream reactor(s) to complete the desired degree of hydrotreatment.

Initially, a set of experiments is conducted under typical hydrotreatment conditions (where the Liquid Hourly Space Velocity (LHSV) is varied, for example, from 0.1 to 10 1/h and the temperature is varied, for example, from 200° C. to 400° C.) to accomplish hydrodemetallation of the metal(s) contained in the feedstock and to concentrate the deposition of the metals in the most upstream reactor(s).

Under such suitable operating conditions, the metals can be deposited in a relatively narrow band in the most upstream reactor(s) of a multi-stage reactor system. Once the upstream catalyst loses its secondary reactivity for metals removal, it can be removed or bypassed in the process reactor train and the reactor reloaded with fresh catalyst. This reactor can now be placed downstream in the reactor train (preferably at the most downstream position), to provide the primary catalytic reaction to complete the desired degree of hydrotreatment. Cycling the reactors in this manner extends the overall utilization of the catalyst and can reduce or eliminate equipment downtime.

General Materials, Equipment, and Test Methods

Feedstock

The biochemically-derived PDO is from E.I. du Pont de Nemours and Company. Four different batches of the biochemical PDO made at a pilot plant in Decatur, Ill., were used in the examples below. These batches were 99.4 to 99.8% pure PDO with over 60 unidentified impurities comprising from about 0.2 to about 0.6% of the crude PDO and had an initial UV absorption at 270 nm (20% solution in water (v/v)) varying from 0.61 to 1.83 and a sulfur content varying from 1.3 to 16 ppm.

Catalyst

The catalyst was a commercially available, supported nickel material, C-28-1-01-RS-CDS catalyst (Süd-Chemie Inc., Louisville, Ky.). It is a reduced and stabilized high nickel-content catalyst containing nominally 52% Ni on silica/alumina. It is an extrudate of 1.6 mm size with a surface area of about 250-350 m2/g. The fresh catalyst contains about 200 ppm sulfur.

Reactor

The laboratory reactor is a jacketed stainless steel tube of 17.3 mm inside diameter packed with either 129 or 250 mm height of catalyst. The reactor was heated by hot oil flowing through the reactor jacket. Both PDO and hydrogen entered at the bottom of the reactor and the flow direction was upflow.

Test Method 1. UV absorption

The PDO color quality was measured by a UV/VIS spectrophotometer. Specifically, the broad UV absorption peak at around 270-280 nm correlates strongly with the presence of color precursors in the PDO and color in the polymers made therefrom. Hydrogenation converts the color precursors and color compounds, reducing the UV-270 nm absorption. Therefore, absorption at UV-270 nm is used as a measure of the extent of hydrogenation. All the UV analyses were measured using a HP 8453 UV/VIS (Hewlett-Packard, Palo Alto, Calif.) spectrophotometer after diluting the PDO to a 20% concentration by volume with water. The results are reported in the Examples at this 20% dilution.

Test Method 2. Sulfur Analysis

The sulfur was analyzed by a Perkin-Elmer 3300RL Inductively Coupled Plasma (ICP) analyzer. Liquid samples were analyzed by direct injection into the analyzer. Catalyst samples were dissolved in acids and then analyzed as aqueous solution. Glossary AU Absorption Unit kPa Kilo Pascal MPa Mega Pascal LHSV Liquid Hourly Space Velocity, 1/h ppm Part per million of weight scc/g Standard cubic centimeters per gram ° C. Degree Celsius mm millimeter g gram h hour approx approximately

EXAMPLES Example 1

Run-36 was conducted in a single reactor with 250 mm catalyst packing at various temperatures (80° C., 100° C., and 120° C., but mostly at 100° C.) at 2860 kPa, 0.8 1/h LHSV with H2 to PDO flow ratio of 6.1 scc/g. The feed had 16 ppm sulfur and the run continued until the catalyst was significantly deactivated. At the end of the run, the catalyst was taken out in segments and analyzed for its sulfur content.

FIG. 2 shows the sulfur profile in the reactor, indicating that sulfur deposition is predominantly near the reactor entrance.

Furthermore, this test showed that the desulfurization rate is much faster than the hydrogenation rate to remove color. The following Table 1 shows that at various conditions of operation while the percentage reduction in UV-270 nm varies from 65 to 94%, the sulfur concentration in the PDO exiting the reactor is below the detection limit of 1 ppm. TABLE 1 Temperature LHSV % Reduction % Reduction ° C. 1/h UV-270 Sulfur 80 0.8 65.3 approx. 100 100 0.4 91.4 approx. 100 100 0.8 87.4 approx. 100 100 1.2 78.1 approx. 100 120 0.8 94.2 approx. 100

Example 2

Run-37 was conducted in a single reactor with 250 mm catalyst packing at various temperatures (80° C., 100° C., and 120° C., but mostly at 100° C.) at 2860 kPa, 0.8 1/h LHSV with H2/PDO=6.1 scc/g. The run continued until the reactor bed was partially deactivated but was still able to meet the desired UV specifications. The run was stopped and the catalyst was removed in segments and analyzed for sulfur.

FIG. 2 shows more distinctly that sulfur deposits predominantly near the entrance of the bed. The partially used bed (Run-37) shows most of the sulfur accumulated in the front one third of the bed. The extensively used catalyst bed (Run-36) shows most of the sulfur accumulated in the first half of the bed.

Example 3

Three segments of the spent catalyst from Run-36 were sampled to represent poisoned catalysts with different levels of sulfur on the catalyst. The hydrogenation activities of these samples and that of a sample of the fresh catalyst were measured under identical conditions of 3.8 1/h LHSV, 100° C., and 2860 kPa.

FIG. 3 shows that the catalyst activity for color removal, as measured by percent reduction in the absorption at UV-270 nm, decreases with increasing level of sulfur accumulated on the catalyst.

Example 4

Run-46 and Run-47 were conducted at 80° C. and 120° C., holding all other operating parameters, including the run duration identical (1.2 1/h LHSV, 2860 kPa, H2/PDO of 7.8 scc/g, for a duration corresponding to about 380 kg of PDO/kg of catalyst).

FIG. 4 shows the distribution of sulfur in the bed. At higher temperatures the rate of desulfurization increased, depositing the sulfur closer to the reactor entrance. This example demonstrates that by selecting the proper temperature, the sulfur deposition profile in the bed can be changed to accumulate the sulfur in a selected zone of the bed, thus protecting the entire bed from deactivation.

Example 5

Run-70 was carried out in two reactors arranged in series. The first reactor was packed with a portion of the poisoned catalyst from Run-36, which was well deactivated and had 6700 ppm of sulfur. This reactor served as a guard bed for the second reactor. The second reactor was packed with fresh catalyst containing 330 ppm sulfur. The run was carried out at 100° C., 1.2 1/h LHSV, 2860 kPa, H2/PDO of 7.8 scc/g for a duration corresponding to 1000 kg of PDO/kg catalyst.

Analysis of the poisoned catalyst in the guard bed reactor showed significant additional sulfur removal. The sulfur profile of the downstream reactor, which was protected by the guard bed, showed substantially lower sulfur deposition, as shown in FIG. 5.

Example 6

In order to evaluate the effect of guard bed in protecting the downstream bed, Run-71 was carried out in a single reactor, identical to the second reactor of Example 5, except that no guard bed was placed upstream of this reactor. The run conditions were identical to those of Run-70.

FIG. 5 shows the effect of the guard bed in protecting the fresh catalyst. In this example, the guard bed reduced the average sulfur deposition on the catalyst from 3000 ppm to about 1350 ppm, or by 55%.

Example 7

Run-76 was made with two reactors in series, both loaded with fresh catalyst, with a combined LHSV of 1.2 1/h at 120° C., 2860 kPa, and 6.9 scc of H2/g of PDO for a duration corresponding to 2100 kg PDO/kg catalyst. During this period the UV-270 of the product increased to 0.1 Absorption Units (AU). After completion of the run, the catalyst in the upstream reactor was replaced with fresh catalyst and this reactor was placed downstream of the reactor originally positioned downstream. The run with spent catalyst in the first reactor, Run-77, was carried out under conditions identical to those in Run-76 to a point where the UV-270 of the product reached the same level (0.1 AU) as in Run-76. This duration corresponded to 1600 kg PDO/kg catalyst. Since this was achieved by replacing only one reactor with fresh catalyst, two catalyst change outs would have given 3200 kg PDO/kg catalyst utilization. Thus compared with Run-76, it presented about 52% increased utilization of catalyst.

The exact percent increase in catalyst utilization may vary somewhat with the targeted product quality and operating conditions. This example clearly showed that application of the invention results in a significant increase in overall utilization of the catalyst relative to that achievable without the invention.

Example 8 Prophetic

For the hydrotreatment of petroleum feedstocks containing organometallic poisons, the following example is proposed. A set of experiments is conducted under typical hydrotreatment conditions (where the LHSV is varied, for example, from 0.01 to 10 1/h and the temperature is varied, for example, from 200° C. to 400° C.) to accomplish hydrodemetallation of the metal(s) contained in the feedstock and to concentrate the deposition of the metals in the most upstream reactor(s).

Under such suitable operating conditions, the metals can be deposited in a relatively narrow band in the most upstream reactor(s) of a multi-stage reactor system. Once the catalyst in the upstream loses its secondary reactivity for the removal of metals, the reactor can be removed or bypassed in the process reactor train and the reactor is then reloaded with fresh catalyst. This reloaded reactor is now placed downstream in the reactor train (preferably at the most downstream position), to provide the primary catalytic reaction to complete the desired degree of hydrotreatment. Cycling the reactors in this manner extends the overall utilization of the catalyst. 

1. A method for extending catalyst utilization comprising: (a) passing a feedstock and hydrogen through at least two serially-connected reactors in positions R₁, R₂, . . . . R_(n), wherein n is the number of reactors, each reactor containing a catalyst, for a period until the catalyst in at least one most upstream reactor is deactivated for a secondary catalytic reaction or until the product from the most downstream reactor fails to meet a desired specification; (b) bypassing the at least one most upstream reactor containing the deactivated catalyst of step (a) to pass the feedstock and hydrogen into at least one downstream reactor; (c) reloading the at least one bypassed reactor of step (b) with fresh catalyst; (d) placing the at least one reloaded reactor of step (c) downstream of at least one of the serially-connected reactors that were not reloaded with fresh catalyst in step (c); and (e) repeating steps (a) through (d) to meet the product specification.
 2. The method of claim 1, wherein the catalyst is selected from the group consisting of a zero-valent element of one or more of the Group VIII elements of the Periodic Table.
 3. The method of claim 2, wherein the catalyst is selected from the group consisting of Fe, Co, Ni, Ru, Rh, Pd, Os, Ir, and Pt.
 4. The method of claim 2, wherein the catalyst is a porous metal structure.
 5. The method of claim 4, wherein the catalyst comprises Raney® nickel.
 6. The method of claim 2, wherein the catalyst is supported on a support.
 7. The method of claim 6, wherein the support is selected from the group consisting of carbon, alumina, silica, titania, silica-alumina, silica-titania, titania-alumina, clays, aluminosilicates, zeolites, water insoluble salts of calcium, water insoluble salts of barium, water insoluble salts of strontium, compounds thereof and combinations thereof.
 8. The method of claim 7, wherein the catalyst comprises nickel with heavy nickel loading supported on extrudates of silica/alumina.
 9. The method of claim 1, wherein the pressure is from about 100 kPa to about 20 MPa.
 10. The method of claim 1, wherein the hydrogen is fed to one or more of the reactors.
 11. The method of claim 1, wherein the temperature is from about minus 50° C. to about 400° C.
 12. The method of claim 1, further comprising adjusting the temperature to compensate for the number of serially-connected reactors in each of steps (b) and (d) for obtaining product with the desired specifications.
 13. The method of claim 1, wherein the feedstock is an organic liquid comprising a compound that can be hydrogenated by a catalyst of one or more of the Group VIII elements of the Periodic Table.
 14. The method of claim 1, wherein in step (d) the at least one reloaded reactor is placed downstream of the most downstream reactor that was not reloaded with the fresh catalyst in step (c).
 15. A method for extending catalyst utilization comprising: (a) passing an organic feedstock and hydrogen through at least two serially-connected reactors in positions, R₁, R₂, . . . R_(n), wherein n is the number of reactors, each reactor containing a porous metal or supported catalyst, for a period until the catalyst in at least one most upstream reactor is deactivated for desulfurization or until the product from the most downstream reactor fails to meet a desired specification; (b) bypassing the at least one most upstream deactivated reactor of step (a) to pass the feedstock and the hydrogen into downstream reactors; (c) reloading the at least one bypassed reactor of step (b) with fresh catalyst; (d) placing the at least one reloaded reactor of step (c) downstream; and (e) repeating steps (a) through (d) to meet a desired specification.
 16. The method of claim 15, wherein the temperature is from about minus 50° C. to about 400° C.
 17. The method of claim 15, wherein the pressure is from about 100 kPa to about 20 MPa.
 18. The method of claim 15, further comprising adjusting the temperature to compensate for the number of serially-connected reactors in each of steps (b) and (d) to meet the desired product specifications.
 19. The method of claim 15, wherein the feedstock comprises 1,3-propanediol.
 20. The method of claim 15, wherein in step (d) the at least one reloaded reactor is placed last in the reactor series, in position Rn.
 21. The method of claim 15, wherein the catalyst comprises nickel with heavy nickel loading supported on extrudates of silica/alumina.
 22. A method for extending catalyst utilization comprising: (a) passing a biochemically-derived organic feedstock containing 1,3-propanediol and hydrogen through at least two serially-connected reactors in positions R₁, R₂, . . . R_(n), wherein n is the number of reactors, each reactor containing a catalyst comprising nickel with heavy nickel loading supported on extrudates of silica/alumina, for a period until the catalyst in the most upstream reactor is deactivated for desulfurization or until the product from the most downstream reactor fails to meet a desired specification; (b) bypassing the reactor in position R₁ to pass the feedstock and the hydrogen into the reactor in position R₂; (c) reloading the bypassed reactor of step (b) with fresh catalyst; (d) placing the reloaded reactor of step (c) in the reactor series downstream at position Rn; and (e) repeating steps (a) through (d) to meet a desired specification.
 23. The method of claim 22 further comprising adjusting the temperature to compensate for the number of serially-connected reactors in each of steps (b) and (d).
 24. The method of claim 22, wherein the temperature is from about minus 50° C. to about 200° C.
 25. The method of claim 22, wherein the temperature is from about 80° C. to about 140° C. 